The present invention relates to a process, reactor system and plant for the production of methanol. In particular, it relates to a process, reactor system and plant for producing methanol from hydrogen and carbon oxides.
Methanol is synthesised in large volumes annually by the conversion of a carbonaceous feedstock, such as natural gas, into a mixture of carbon oxides and hydrogen. Such a mixture of gases in often referred to as synthesis gas.
The conversion of a hydrocarbon-containing feedstock, such as natural gas, into synthesis gas can be achieved by steam reforming, by partial oxidation, or by a combination of these processes.
In steam reforming a mixture of desulphurised hydrocarbon feedstock, such as natural gas, and steam is passed at high temperature, typically at a temperature of from about 600° C. to about 1000° C., and elevated pressure, typically from about 10 bar up to about 50 bar, over a suitable reforming catalyst, such as a supported nickel catalyst. One commercially recommended catalyst which is suitable for this purpose uses a mixture of calcium and aluminum oxides as support for the nickel. The principal reaction is:CH4+H2O≈CO+H2.
The reaction products themselves are further subject to the reversible “water gas shift” reaction in which carbon dioxide and hydrogen are produced from carbon monoxide and steam:CO+H2O≈CO2+H2.
Another method for producing synthesis gas involves the use, wholly or in part, depending upon the carbon to hydrogen ratio in the hydrocarbonaceous feedstock, of direct catalytic or non-catalytic partial oxidation or secondary/autothermal reforming with oxygen. In the case of methane this occurs according to the following equation:CH4+½O2≈CO+H2.
A combination of steam reforming and partial oxidation or secondary/autothermal reforming can also be used.
Conversion of the carbon oxides and hydrogen to methanol occurs according to the following reactions:CO+2H2≈CH3OHCO2+3H2≈CH3OH+H2O.
These reactions are conventionally carried out by contacting the synthesis gas with a suitable methanol synthesis catalyst under an elevated synthesis gas pressure, typically in the range of from about 50 bar up to about 100 bar, usually about 80 bar, and at an elevated methanol synthesis temperature, typically from about 210° C. to about 270° C. or higher, e.g. up to about 300° C.
Suitable methanol synthesis catalysts include copper containing catalysts with a catalyst comprising a reduced zinc oxide/copper oxide mixture being particularly-suitable.
As with many reactions it is desirable to achieve the maximum rate of reaction per weight of catalyst or per volume of the reactor.
A conventional methanol synthesis plant can be considered to comprise four distinct parts, namely:    1. a reforming plant, which produces a mixture of carbon oxides and hydrogen from a hydrocarbon feedstock;    2. a compression stage which lifts the carbon oxides and hydrogen mixture to a higher pressure suitable for downstream methanol synthesis;    3. a methanol synthesis section, in which crude methanol is produced from the carbon oxides and hydrogen; and    4. a distillation section, in which the final refined methanol product is produced from the crude methanol.
A number of different types of reformer for use in part 1 of the methanol synthesis plant i.e. the reforming plant, are known in the art. One such type is known as a “compact reformer” and is described in WO-A-94/29013, which is incorporated herein by reference and which discloses a compact endothermic reaction apparatus in which a plurality of metallic reaction tubes are close-packed inside a reformer vessel. Fuel is burned inside the vessel, which comprises air and fuel distribution means to avoid excessive localised heating of the reaction tubes. In a compact reformer of this type heat is transferred from the flow gas vent and from the reformed gas vent of the reformer to incoming feedstock, fuel and combustion air. Other types of reformer are not as efficient as the compact reformer in transferring heat internally in this way. However, many other reformer designs are known and some are described in EP-A-0033128, U.S. Pat. No. 3,531,263, U.S. Pat. No. 3,215,502, U.S. Pat. No. 3,909,299, U.S. Pat. No. 4,098,588, U.S. Pat. No. 4,692,306, U.S. Pat. No. 4,861,348, U.S. Pat. No. 4,849,187, U.S. Pat. No. 49,090,808, U.S. Pat. No. 4,423,022, U.S. Pat. No. 5,106,590 and U.S. Pat. No. 5,264,008, U.S. Pat. No. 5,264,008 and WO 98/28071 which are incorporated herein by reference.
In a conventional plant, synthesis gas is compressed in passage from the reforming plant to the methanol synthesis zone. This compression stage is generally present in order to provide the required pressure of from 50 bar to 100 bar in the methanol synthesis zone. The compressed gas is then passed to the methanol synthesis section.
In U.S. Pat. No. 4,594,227 apparatus for carrying out a catalytic chemical reaction is described which comprises a vertical, annular, intercylinder space which is divided by radially extending vertical partition walls into a plurality of chambers some of which include heat-exchanging tubes. Segments containing no heat-exchanging tubes may be packed with catalyst and utilised adiabatically to preheat the reaction gases. In use, the reaction gases will pass outwardly through this first segment, where any reaction will cause heating, they then are transmitted through the annular space surrounding the intercylinder space before travelling inwardly through the segment containing catalyst and cooling tubes where further reaction will occur.
Whilst this arrangement does offer certain advantages, it also suffers from various drawbacks. A principle disadvantage arises from the multi-segmental radial flow. This flow pattern causes the gas velocity to vary as the flow traverses from the centre of the reactor to the outside and back, due to the changing cross-sectional area of the segments. This changing velocities of the segments causes the heat transfer coefficient between the reacting gases and the cooling medium in the tubes to vary. In particular the heat transfer will increase as the gas velocity increases and will decrease as the gas velocity is reduced.
Thus the multi-segmental arrangement of the radial flow apparatus in U.S. Pat. No. 4,594,227 does not allow the gas velocity pattern and resultant heat transfer pattern to be optimised.
Various methanol production processes are known in the art, and reference may be made, for example, to U.S. Pat. No. 5,610,202, U.S. Pat. No. 4,968,722, U.S. Pat. No. 5,472,986, U.S. Pat. No. 4,181,675, U.S. Pat. No. 5,063,250, U.S. Pat. No. 4,529,738, U.S. Pat. No. 4,595,701, U.S. Pat. No. 5,063,250, U.S. Pat. No. 5,523,326, U.S. Pat. No. 3,186,145, U.S. Pat. No. 344,002, U.S. Pat. No. 3,598,527, U.S. Pat. No. 3,940,428, U.S. Pat. No. 3,950,369, WO-A-98/28248 and U.S. Pat. No. 4,051,300 which are incorporated herein by reference.
Various suggestions have been made for modifications to the plant design with a view to improving the economics of the production process.
Several suggestions for improving the efficiency of the reaction have been made which incorporate the use of multiple reaction stages. For example, in U.S. Pat. No. 5,631,302 it is suggested that the methanol synthesis section should include two separate synthesis reactors. In this arrangement, the synthesis gas is passed to the first synthesis reactor, which is a shaft reactor containing a fixed bed of a copper-containing catalyst. The reaction in this shaft reactor is carried out adiabatically and in the absence of any recycling of synthesis gas. The product stream from this first reactor, which contains methanol vapour, is cooled to condense the methanol which is separated from the unreacted gaseous. components of the first product stream. These unreacted gaseous components are then compressed, heated and fed to the second reactor where they react to form methanol. The second reactor is preferably a tubular reactor in which the copper catalyst is indirectly cooled by water which is boiling under high pressure. The product stream from the second reactor is cooled and the methanol is removed by separation. Any unreacted gaseous components are compressed and heated before being returned to the second reactor.
Thus in U.S. Pat. No. 5,631,302 the first reactor is located outside the main reactor loop and simply serves to modify the composition of the feed gas before it enters the main reaction loop. The arrangement of U.S. Pat. No. 5,631,302 is said to be useful where the synthesis gas feed has a CO2:CO ratio which exceeds 2:1.
An alternative arrangement is suggested in U.S. Pat. No. 5,827,901. In this arrangement two synthesis reactors are connected in series such that the product stream from the first reactor is passed directly to the inlet of the second reactor. The first reactor is a water cooled reactor in which the catalyst is located in tubes through which the gaseous reactants flow. The second reactor may be selected from a variety of designs. Whichever design is used, cooling in the second reactor is provided by counter-current heat exchange with the feed synthesis gas before it is fed to the first reactor.
This arrangement allows for cooler exit temperatures from the second reactor to be achieved than are conventionally achievable. However, whilst the lower temperature may allow the reaction equilibrium to move towards completion, it may also reduce the rate of reaction, and therefore may require more catalyst per unit of product.
Other examples include U.S. Pat. No. 5,427,760 in which two reaction stages are used in an attempt to achieve a higher overall conversion to the desired ammonia than can be achieved in a single stage and U.S. Pat. No. 4,867,959 in which two or more reaction stages are described, with cooling between each stage, to increase conversion. As discussed by Kobayashi and Green in a paper presented to the 1990 World Methanol Conference, this approach can be extended to include a large number of stages. This paper also illustrates the optimum rate line for methanol synthesis.
Whilst an optimum rate line is known, a near-optimum reaction profile is not practical in commercial arrangements. This is because such a profile would generally require the reaction to start at high temperature and gradually fall as the reaction proceeds. Some suggestions have been made to produce a system which approaches the optimum rate line, such as those in the Kobayashi and Green paper however, a commercially satisfactory arrangement has not been realised.
Thus it will be understood that whilst the systems of the prior art go some way to addressing the problems associated with reducing the operating and/or investment costs of producing methanol, various disadvantages and drawbacks remain and there is still a requirement for alternative arrangements which will address at least some of these problems.